Recycle hydrocracking process for converting heavy oils to middle distillates

ABSTRACT

HEAVEY MINERAL OILS BOILING MAINLY IN THE 700-1000*F. RANGE ARE SELECTIVELY CONVERTED TO MIDDLE DISTILLATED BOILING IN THE 300-700*F. RANGE WITH A MINIMUM PRODUCTION OF LIGHTER HYDROCARBONS, BY CONTACTING THE FEED UNDER HYDROCRACKING CONDITIONS FIRST WITH A RELATIVELY NONACIDIC, HYDROFINING TYPE CATALYST SUPPORTED ON A SUBSTANTIALLY NONCRACKIG REFRACTORY OXIDE CARRIER, AND THEN WITH AN ACIDIC, ZEOLITE HYDROCRACKING CATALYST. UNCONVERTED OIL FROM THE SECOND CONTACTING ZONE IS RECYCLED TO THE FIRST ZONE. MOST OF THE DESIRED CONVERSION IS EFFECTED IN THE FIRST CONTACTING ZONE, WHILE THE SECOND ZONE IS UTILIZED TO EFFECT A MINOR ADDITIONAL CONVERSION TO DESIRED PRODUCT, AND TO HYDROCRACK (AND THUS PREVENT THEIR BUILDUP IN THE RECYCLE OIL) HEAVY, FUSED-RING POLYCYCLIC AROMATICS SUCH AS COROMENES AND OVALENES.

Jan'. 12,1971 y|r. :.w' 3o|: ErAL n 1 3,554,898 RECYCLE HYDROCRAGKING PROCESS FOR CONVERTING I -l i I HEAVY OILS TO MIDDLE DISTILLATES FilledV Aug. 29, 1968 NH3 f/F afs/R50) Arranz/Ey United States Patent Office U.S. Cl. 208-59 Claims ABSTRACT F THE .DISCLOSURE Heavy mineral oils boiling mainly in the 700-1000 F. range are selectively converted to middle distillates boiling in the 300-700 F. range with a minimum production of lighter hydrocarbons, by contacting the feed under hydrocracking conditions -rst with a relatively nonacidic, hydroning type catalyst supported on a substantially noncracking refractory oxide carrier, and then with an acidic, zeolite hydrocracking catalyst. Unconverted oil from the second contacting zone is recycled to the irst zone. Most of the desired conversion is effected in the rst contacting zone, while the second zone is utilized to effect a minor additional conversion to desired product, and to hydrocrack (and thus prevent their buildup in the recycle oil) heavy, fused-ring polycyclic aromatics such as coronenes and ovalenes.

BACKGROUND AND SUMMARY OF THE INVENTION Recent years have witnessed a phenomenal growth in the development and application of catalytic hydrocracking processes. By far the greater part of these efforts have been directed toward methods for converting gas oils to products boiling in the gasoline range. The best catalysts developed to date for this purpose are those comprising a highly active cracking base, eg., of the crystalline zeolite type, combined with a highly active hydrogenation component such as palladium or platinum. These catalysts are very efficient for converting middle distillate oils boiling in the 40G-750 F. range to gasoline. Due to the extensive use, and prospective use, of middle distillate stocks as feed to these gasoline-producing hydrocracking units,

. and to other economics, geographic and seasonal factors,

a need is being felt in the industry to provide additional middle distillate stocks to meet the demand for other products such as turbine fuels and diesel fuels. An obviously desirable source of such additional middle distillate stocks comprises the heavy distillates boiling above about 700 P., which have heretofore been diverted largely to fuel oils because of the lack of an economical method for converting them to lower boiling products.

In the initial investigation of hydrocracking techniques for converting 700 F.-| hydrocarbon feeds to lower boiling materials, it became apparent that deep hydrocracking to produce gasoline in a single conversion step was impractical, rstly because inordinate amounts of butanes and dry gas were produced, and secondly because the gasoline product was always of extremely low octane quality. It become apparent therefore that for purposes Patented Jan. 12, 1971 of gasoline production, hydrocracking techniques were of primary value in connection with middle distillate feedstocks, and that the hydrocracking of higher boiling stocks would be feasible only if conversion to gasoline could be minimized and the production of middle distillate maximized. A primary objective in the hydrocracking of these heavy feeds thus became to obtain maximum middle distillate/ gasoline ratios in the resulting product.

It was then found that the catalysts most useful for converting middle distillates to gasoline were least useful for converting heavy feedstocks selectively to middle distillates, in that large amounts of dry gases, butanes and gasoline were produced. A great many conventional hydrocracking catalysts were tested in an attempt to find one which would convert such feedstocks more selectively to middle distillate products. In all cases it was found that the desired selectivity could be maintained only by operating at low overall conversion rates, entailing prohibitively high recycle rates. But, on testing certain catalysts not conventionally regarded as hydrocracking catalysts, substantially more promising results were obtained. Specifically, it was found that by using a type of catalyst commonly employed for catalytic hydroining, composed of nickel and molybdenum oxides and/or suldes supported on activated alumina, the selectivity of conversion to middle distillate products was excellent even at relatively high conversions of e.g. 40-6() volume-percent per pass. Moreover, for feedstocks containing organic nitrogen, it was found that these catalysts were not only more selective, but more active (on the basis of temperature required for a given total conversion) than even the most active hydrocracking catalysts based on zeolite cracking bases.

But in attempting to utilize these nickel-molybdenumalurnina catalysts for total conversion of the heavy feedstocks to middle distillates, i.e., with total recycle of unconverted oil, a problem was encountered of maintaining stable hydrocracking temperatures permitting commercially feasible run lengths. Analysis of the recycle oil showed a progressive building therein of heavy polycyclic, condensed ring aromatics of the coronene and ovalene types, which are known coke precursors and thus catalyst deactivators. It should be understood that these coronene-ovalene contaminants are not native constituents of crude oils, or any of the virgin distillates therefrom. Neither are they found in detectable quantities in the unconverted oils from conventional catalytic cracking, coking, thermal cracking, or the like.. The presently available evidence indicates that, in the present case, such compounds are synthesized in the first catalytic conversion zone, as a result of the relatively high temperatures employed and the relatively low hydrogenation and cracking activity of the hydrofining-type catalyst employed. Their solubility in hydrocarbon oils is very low, amounting to only about a few parts per million, up to perhaps parts per million of the more soluble species. These characteristics lead to two diiculties. Firstly, as indicated above they tend to coke and deactivate the catalyst as a result of the high temperatures employed. Secondly, with total recycle of unconverted oil, their concentration builds up beyond the solubility limits, after which they begin to plate out in cooler portions of the system, particularly heat exchange surfaces, resulting in plugging problems and reduced heat exchange efliciency. Examples of such compounds which have been found in solid deposits removed from heat exchangers are as follows:

ooo ooo OZ) (#522) en Tribenzcoronenes OOO :ses

Benzovalenes (-UZ) Tribenzcoronenes (-saz) Tetrabenzcoronenes It will be apparent that in many instances, the above structural formulae represent merely one isomer of the given generic name. The Z factors noted refer to the hydrogen deficiency of the respective hydrocarbons, from the general formula, CnH2n+z.

The present invention is based essentially upon our discovery that the acidic, zeolite-type hydrocracking catalysts (which normally are undesirable for selective conversion to middle distillate products) can be advantageously employed herein for hydrocracking the coronene-ovalene compounds contained in the effluent from a conversion zone utilizing nonacidic hydrofining type catalysts, while at the same time effecting a limited additional conversion of the heavy feed to middle distillate products with a minimum production of gasoline. The effluent from the zeolite hydrocracking zone can then be fractionated to recover desired middle distillate products, and

a heavy recycle oil of reduced coronene-ovalene content which can be recycled continuously to the rst hydrocracking zone employing the nonacidic hydroning type catalyst for additional selective conversion to middle distillate product. This .unique application of the zeolite type hydrocracking catalysts requires firstly that relatively low overall conversion rates be maintained therein in order to minimize the production of gasoline and lighter products. But secondly, it is also necessary that relatively high temperatures, above about 600 F., be maintained, for at lower temperatures it has been found that such catalysts are relatively ineffective for converting coroneneovalene compounds. These two seemingly inconsistent requirements are reconciled herein by maintaining a sulficient ammonia concentration in the zeolite conversion zone to repress overall hydrocracking activity suiciently to permit temperatures to be raised to adequate coroneneovalene conversion levels without encountering excessive hydrocracking to gasoline and lighter products. Fortuitously, in most cases sufficient ammonia is present in the efuent from the first conversion zone to achieve this objective, but in some cases it may be desirable to add extraneous ammonia to the zeolite hydrocracking zone.

Maintaining ammonia in the zeolite hydrocracking zone is also advantageous from the standpoint of increasing the selectivity of hydrocracking to middle distillate products. Ammonia-moderated, high temperature hydrocracking appears to increase the relative activity of the zeolite catalysts for hydrocracking some of the heavier feed components (including coronene-ovalene compounds) as opposed to the lighter, middle distillate components. This may possibly be explained on the basis of differing basicity of the light and heavy aromatic feed components; polycyclic aromatics are more basic than the monocyclic components, and hence can more strongly compete with ammonia for the acidic cracking centers on the catalyst.

The following process description will illustrate more in detail the contemplated practical application of the above principles and discoveries.

DETAILED PROCESS DESCRIPTION Reference is made to the accompanying drawing which is a simplified flow diagram illustrating a preferred modification of the invention. Initial raw feed is brought in through line 2, mixed with heavy recycle oil from line 4 and with recycle and make-up hydrogen from lines 6 and 8, and the resulting mixture is then passed via preheater 10 into hydrofiner-hydrocracker 12, containing one or more beds of a conventional hydroiining catalyst to be described hereinafter. In hydroner-hydrocracker 12, process conditions are adjusted so as to achieve substantial desulfurization, dentrogenation, and conversion of the feedstock to lower boiling material. To achieve these objectives, process conditions are adjusted within the following general ranges:

HYDROFININ G-HYDROCRACKING CONDITIONS Broad Preferred range rango Temperature, F. 725-875 750850 Pressure, p.s.i.g 1, OOO-5, 000 1, 50H3, 000 LHSV 0. 5-10 1-5 Hg/Oil, MSCF/B 4-20 6-12 about 2/1. For this purpose, preferred conversion levels (based on true boiling points) normally lie in the range of about 35-55 percent to products boiling below the true 5 percent boiling point of the feed.

In achieving the foregoing conversion levels, it will normally be found that the feedstock will be at least about 95 percent desulfurized, and at least about 90 percent denitrogenated. All of these reactions are exothermic, and hence it may be desirable to inject cool quench hydrogen at one or more points in the reactor, as illustrated via line 14.

Effluent from reactor 12, containing ammonia and hydrogen sulfide generated therein, is withdrawn via line 16, blended with additional ammonia from line 18 if desired, and with makeup' recycle hydrogen from line 20. The resulting mixture is then passed into zeolite catalyst hydrocracker 22 via heat exchanger 24, which normally functions to reduce the temperature of the mixture about 20- 80 F. below the temperature of the effluent from reactor 12. The degree of temperature reduction depends primarily upon the amount of ammonia present in the transferred mixture. VThe effective cracking activity of the catalyst in reactor 22 is a function primarily of the ammonia partial pressure therein, and as stated above, it is necessary to reduce catalyst activity sufficiently to permit the use of temperatures in reactor 22 which, in the absence of ammonia, would bring about undesirable, nonselective hydrocracking, but which temperatures are desired herein in order to achieve adequate conversion of the coronene-ovalene compounds. Contemplated operating conditions to be maintained in reactor 22 are illustrated as follows:

ZEOLITE CATALYST HYDROCRACKING CONDITONS Broad Preferred range range Temperature, F 625-825 70D-800 Pressure, p.s.i.g 1, OOO-5, 000 1, 50G-3, 000 Ll-ISV 1-20 2-10 lf2/Oil, MSCF/B 4-20 6-12 sion levels it is found that the ratio of middle distillates/ l 'gasoline synthesized -in reactor 22 may be in the same high range as that prevailing in the effluent from reactor 12. At the same time, sufficient conversion, normally between about and 90 percent, of the coronene-ovalene content of the feed entering reactor 22 takes place to prevent the buildup of such materials in the unconverted oil which will be recycled to reactor 12. The space velocity in reactor 22 is normally substantially higher, e.g., two to four times higher, than the space velocity in reactor |12. This is lbecause the zeolite catalyst in reactor 22 is not being used at maximum efficiency (i.e., maximum conversions), and hence from economic considerations it is desirable to use only as much of such catalyst as is required to convert the coronene-ovalene compounds, and to carry out most of the desired selective conversion to middle distillate in reactor 12 where the catalyst is more specifically tailored for that purpose.

It will be apparent from the foregoing that the ammonia partial pressure to be maintained in reactor 22 will vary in direct proportion to the average bed temperature, assuming the same intrinsic catalyst activity. At high temperatures within the indicated ranges, the ammonia partial pressure may amount to as much as 150 p.s.i. or higher in order to limit conversion to the stated ranges. At the lower temperatures however, the ammonia partial pressure may be as low as 0.1 p.s.i.

Effluent from hydrocracker 22 is withdrawn via line 26, cooled and condensed in heat exchanger 28, blended with wash water injected via line 30, and then transferred to high-pressure separator 32, from which recycle hydrogen is withdrawn via line 34. Spent wash water containing dissolved ammonia and some hydrogen sulfide is withdrawn via line 36, while the remaining high-pressure hydrocarbon condensate `is ashed via line 38 into low-pressure separator 40, from which light flash gases are exhausted via line 42. Low-pressure condensate in separator 40 is then transferred via line 44 to fractionating column 46 from which the minor gasoline product is -withdrawn overhead via line `48, and the major middle distillate product via side-cut line 50. The remaining unconverted oil, boiling mainly above 700 F., is withdrawn as bottoms via line 52 and recycled to fresh feed line 2 as previously described.

Obviously, many modifications of the above processing scheme may be utilized without departing from the scope of the invention. In particular, it should be noted that, instead of employing two reactors 12 and 22, a single reactor may be utilized containing the hydroning catalyst near the inlet end thereof and the zeolite catalyst near the outlet end, with appropriate means for introducing quench hydrogen and/ or ammonia at a level in the reactor between the two catalysts.

FEEDSTOCKS As will be apparent from the foregoing, the process of this invention is designed exclusively for the hydrocracking of mineral oil feedstocks containing a substantial proportion, preferably a major proportion, of hydrocarbons having a true boilingv point above about 700 F., and up to about 1200 F. Specifically preferred feedstocks are those containing less than about 10 Volume-percent of material boiling below 650 F., at least about 70 volume-percent of material boiling between 7 00 F. and 1000 F., and at least about 20 volume-percent of material boiling above 800 F. Feedstocks of this nature cannot be distilled at atmospheric pressure without substantial decomposition. They are normally derived from the vacuum distillation of crude oils, or by the deasphalting of residual oils. The heaviest fractions of catalytic cracking cycle oils, coker distillates and/or thermally cracked oils may also be utilized, either alone or in admixture with the preferred straight-run vacuum distillates or deasphalted residual oils. These feedstocks will normally contain at least about 10 weight-percent, and up to about 70 percent, of aromatic constituents; sulfur in amounts of about 0.01 to 3 percent by weight, and nitrogen in amounts of about 0.001 to 2 percent by weight.

The feedstocks and products described herein are characterized mainly on the basis of boiling range. Unless otherwise stated, when boiling ranges are given, atmospheric boiling points are intended. Since the feedstocks cannot be distilled at atmospheric pressure, the atmospheric boiling ranges cited are -calculated from standard ASTM D-1160 distillations carried out at about 1 millimeter of mercury. The D-l distillation is operated essentially without reflux, and hence does not provide a sharp fractionation. For a more accurate determination of product yields and conversions, a true boiling point (TBP) distillation utilizing reflux is employed in the examples herein. The abbreviation TEBP as applied to the 700 F. end-point product refers to its true end boiling pom DESCRIPTION OF CATALYSTS Suitable hydrofning-hydrocracking catalysts for use in the first contacting zone (reactor 12) include for example the oxides and/or sulfides of molybdenum and/or tungsten, preferably composited with an iron group metal oxide and/or sulfide such as nickel or cobalt. Preferred catalysts comprise sulfided composites of molybdenum oxide and nickel oxide supported on an adsorbent, relatively noncracking mineral oxide carrier such as activated alumina, or any other difiiculty reducible, refractory oxide having a Cat-A activity index below about 25. Preferred catalysts contain about 2-6 weight-percent nickel and 5-25 weightpercent molybdenum, while the preferred carriers comprise activated alumina containing a minor proportion, e.g. 3-25 weight-percent, of coprecipitated silica gel.

The zeolite hydrocracking catalyst utilized in reactor 22 comprises a minor proportion of a transitional metal hydrogenating component (preferably a Group VI-B metal and/or a Group VIII metal, r the oxides or suldes thereof) supported on a major proportion of a zeolite cracking "base, The preferred hydrogenating metals comprise the Group VIII noble metals, especially platinum and/or palladium (in amounts of about 0.1-2 weight-percent), but the Group VI-B metals, particularly molybdenum and/or tungsten may also be utilized, either alone or in combination with a Group VIII metal such as nickel and/or cobalt. The Group VI-B metals are used in amounts of about 5-30 weight-percent, while the iron group metals are preferably employed in amounts of about 1410 weight-percent.

The operative zeolite cracking7 bases comprise crystalline or amorphous aluminosilicates having ion exchange capacities greater than about 1.0 n1eq./ gm., and wherein the zeolitic cations are predominantly hydrogen ions and/ or polyvalent metal ions. The preferred zeolites are crystalline molecular sieves having relatively uniform pore diameters of about 6-14 A. These zeolites may be used as the sole cracking base, or they may :be mixed with one or more nonzeolitic bases such as alumina, or silica-alumina cogel. Suitable crystalline molecular sieves include for example those of the X, Y or L crystal types. These zeolite cracking bases are further characterized by a cracking activity greater than that corresponding to a Cat-A activity index of 40, preferably greater than about 50. The activity index of a catalyst is numerically equal to the volume-percent of gasoline produced in the standard Cat-A activity test as described in National Petroleum News, Aug. 2, 1944, volume 36, page R-537.

A particularly active and useful class of zeolite bases are those having a relatively high SiO2/Al2O3 mole ratio, eg., between about 3 and 10. The preferred zeolites are those having crystal pore diameters between about 8-12 A., and wherein the SiO2/Al203 mole-ratio is between about 3 to 6, including primarily synthetic Y molecular sieve.

Hydrogen and/or decationized zeolites useful herein are more particularly described in U.S. Pat. No. 3,130,006. Suitable polyvalent metal zeolites include for example the magnesium, calcium, zinc and rare earth metal zeolites and the like, or in general any of the polyvalent metal zeolites of Groups I-B through Group VIII. The preferred polyvalent metals are the alkaline earth metals, zinc and the rare earth metals.

The hydrogenating metals may be added to the zeolite cracking base by conventional impregnation methods, and/or by ion-exchange as described in U.S. Pat. No. 3,236,762. It is preferable however to add the Group VI-B metals to the zeolite by means other than ion exchange. For example, a water insoluble compound of the Group VI-B metals may be mixed with the zeolite in the presence or absence of lwater by stirring, mulling, or grinding. A water soluble compound of the Group VI-B metal may be mulled or ground with the zeolite, provided the mixture is relatively dry. Insoluble or undissolved Group VI-B compounds can be formed in the presence of the zeolite., For example, the zeolite can be slurried in a solution of ammonium molybdate or tungstate. Then the slurry is acidied to precipitate molybdic or tungstic acid.

Following one or more the above procedures, the zeolite catalyst may be mixed with one or more relatively nonzeolitic amorphous gels such as alumina or silicaalumina which act as a binder. The resulting powdered mixture is then pelleted and calcined at conventional temperatures of about 900 F. or higher. The resulting catalyst is presulfded if desired.

The following examples are cited to illustrate the invention more specifically, but are not to be construed as limiting in scope:

PREFACE TO EXAMPLES I-V These examples illustrate primarily the improved selectivity of conversion to middle distillate products obtained with the hydroning catalysts employed in the rst stage of this invention, as compared to the zeolite hydrocracking catalysts employed in the second stage. Since there was no recycle of unconverted oil in these examples, the deactivating effects of coronene-ovalene compounds do not appear. In each example, the hydrocracking was carried out at a hydrogen/oil ratio of 7500 s.c.f./barrel, and the feedstock was a straight run vacuum distillate analyzing as follows:

FEED ANALYSIS Gravity, API 19.7 Boiling range, ASTM D-l160 at 1 mm., converted to atmospheric pressure:

Initial B.P., F 555 5% 689 10% 725 50% 841 932 959 Maximum 1016 Volume percent boiling :below 700 F., TBP 6-8 Sulfur, wt.percent 1.54

Nitrogen, wt.percent 0.387 Aromatics, wt.percent 35.2

In these examples, Conversion to 700 F. TEBP Products refers to actual feed disappearance, i.e., minus the volume-percent of feed recovered as product having a true initial boiling point of 700 F. The Diesel/Gasoline Ratio refers to the volume ratio in the product 0f 30G-700 F. boiling range (TBP) diesel fuel to` C5- 300 F. gasoline.

EXAMPLE I The above feedstock was subjected to hydrocracking under two different sets of conditions utilizing a preferred hydrolining type catalyst of this invention consisting of a presulded composite of 3.1 weight-percent nickel oxide and 14.7 weight-percent M003 supported on a coprecipitated alumina-silica cogel containing 3.2 weight-percent SiO2, in the form of J,fw-inch extruded pellets. The principal conditions and results of the runs were as follows:

TABLE 1 Run N0. I-A I-B Catalyst age, hours 56-68 116-140 LHSV, v./v 1.0 0.5 Temperature, "F 790 790 Pressure, p.s.i.g 1, 500 1, 500 Converslon to 700 F. TEBP products, volume percent 40. 6 56. B Diesel/gasoline ratio in product, v./v 6. 1 5. 0 Hydrogen consumption SCF/B. 1, 276 2, 054 CFC; gas make, SCF/B 77.1 94. 3 Butanes, volume/percent oi feed 1. 4 1.7 Diesel fuel product:

Yield, volume/percent of feed 46.6 52. 8 Cetane index 48. 2 47. 2 Nitrogen, ppm 60 The foregoing illustrates a desirable selective conversion to diesel fuel, with a minium production of lighter materials.

EXAMPLE II The operation described in Example I was continued at higher temperatures and pressures to obtain an undesirably high overall conversion, with the following results:

tinued at an elevated pressure of 2500 p.s.i.g., with the following results:

TAB LE 2 TABLE 4 Run No. I-C I-D 5 Run N o. Il-E H-F Catalyst age, hours 364-376 Catalyst age, hours 471-496 543-567 LH V, v. 0.5 LHSV, v./v 0. 5 0.5 Temperature, F 825 850 Temperature, F 790 825 Pressure, p.s.i.g 2, 500 2, 500 Pressure, p.s..g 2,500 2,500 Conversion to 700 F. TEBP products, volum Conversion to 700 F. TEBP products, volume] percen 74. 1 88. 4 percent 58. 4 80. 1 Diesel/gasoline ratlo 1n product v /v 3. 7 2. 1 10 Diesel/gasoline ratio in product, v./v 5. 0 3. 1 (J1-C3 gas make, SCF/B 137 232 Ci-Cg gas make, SCF/B 79. 7 145. 9 Butanes, volume/percent o feed 3. 3 5. 1 Butanes, volume/percent of feed 1. 8 3. 1 Diesel fuel product: Diesel fuel product:

Yield, volume/percent of feed 64. 7 66. 9 Yield, volume/percent of feed 54.5 66. 7 Cetane index 50. 0 47. 6 Oetane index 51. l 50. 8 Nltrogen, p.p.m 10 10 Nitrogen, p.p.m 10 10 The foregoing demonstrates that, even when using the preferred hydroning type catalysts of this invention, the selectivity of conversion to diesel fuel is relatively low at overall conversion levels above about 70 percent. The production of dry gas, butanes and gasolines was much higher than in Example I.

EXAMPLE III The above feedstock was subjected to hydrocracking under several different run conditions using another hydroining type catalyst of this invention comprising a presulfded composite of 4.5 weight-percent Ni and 13 weightpercent M003 supported on an alumina-silica cogel base containing 14 weight-percent of silica, in the form of lAG- inch extruded lpellets. The principal conditions and results were as follows:

The foregoing data shows that good diesel fuel quality can be maintained at temperatures up to 825 F. at a pressure of 2500 p.s.i.g. It is again noted however that the high conversion level of 80.9 percent results in decreased selectivity of conversion to middle distillate product.

EXAMPLE V TABLE 3 Run No. II-A II-B II-O II-D Catalyst age, hours 42-66 90-114 339-351 198-210 LHSV, v./v 1. 0 l. 1.0 1 0 Temperature, F.. 760 7 790 800 Pressure, p.s.i.g 1, 500 1, 500 1, 500 1, 500 Converslon to 700 F percent 33. 0 45. 4 49. 0 56. 6 Diesel] gasoline ratio in product, 9. 1 6. 0 5. 1 4. 7 Hydrogen consumption, SCF/B 662 774 686 935 Cl-Cs gas make, SCF/B 40. 1 73. 7 86. 2 98. 3 Butanes, volume/percent of feed 0. 7 1. 3 1. 3 1. 5 Diesel fuel product:

Yield, volume/percent of feed 31. 7 42.2 41. 0 50. 0 Cetane index 47. 0 45. 0 46. 5 45. 5 Nitrogen, p.p.rn 220 80 60 The above data again illustrates a desirable selective conversion to diesel fuel lproduct with minimum convermagnesium ions. The principal conditions and results were as follows:

TAB LE 5 Run No. III-A III-B III-C Catalyst age, hours- 118-142 379-382 298-310 LHSV, v. 1. o 1.0 1.0 Temperature, F... 760 800 825 Pressure, p.s.i.g 1500 1500 1500 Conversion to 700 percent 25. 9 33. 2 57. 6 Diesel/gasoline ratio in product, 7. 4 4. 3 1. 6 Hydrogen consumption, SCFIB 454 581 1076 Cl-Ca gas make, SCF/B 23. 3 54. 5 134. 4 Butanes, volume/percent of feed. 0. 6 1. 4 5. 5

Diesel fuel product:

Yield, volume/percent of feed 23. 1 27. 1 35. 4 Cetane index 46. 3 44. 3 42. 2 Nitrogen, p.p.rn 2330 230 110 sion to gasoline andv lighter materials. It will be noted however that at the 1500 p.s.i.g. pressure, the quality of the diesel fuel product was relatively low.

EXAMPLE 1V The above data shows that the zeolite catalyst gave relatively poor selectivity of conversion to diesel fuel at conversion levels above about 30 percent. However at conversion levels below 30 percent, the selectivity was commensurate with that obtained in Examples I to IV using Hydrocracking as described in Example III was conhydroiining catalysts. It is hence feasible to use the zeolite type catalysts at low conversion levels in the second stage of the present process, in view of its additional advantage demonstrated in the succeeding example, of effecting conversion of coronene-ovalene compounds.

EXAMPLE VI This example demonstrates that (l) the hydrolining type catalysts employed herein do tend to bring about synthesis of coronene-ovalene compounds when utilized at high temperatures, (2) the zeolite type catalysts do not effectively convert such compounds at low temperatures, but that (3) such zeolite catalysts do effectively convert coronene-ovalene compounds at higher temperatures in the presence of ammonia.

A 120-day hydrocracking run was carried out in a twostage hydrocracking unit wherein the feed was first subjected to catalytic hydrofining, with the total eluent therefrom then being subjected to a rst stage of catalytic hydrocracking employing a zeolite hydrocracking catalyst, and with unconverted oil from the first hydrocracking stage being hydrocracked at lower temperatures in a second zeolite catalyst hydrocracking zone. Unconverted oil from the second hydrocracking zone was continuously recycled thereto.

In this example, the product was gasoline, but the relevant principles in respect to coronene-ovalene synthesis and conversion apply equally in the case of hydrocracking to middle distillate products.

The feed was a blend of about 54 volume-percent of a straight run vacuum gas oil Iand 46 percent of a light catalytic cracking cycle oil analyzing as follows:

Straight Light run catalyti: vacuum cycle gas oil oil Gravity, API 24.0 21.0 ASTM distillation, F.:

IBP 480 456 10% 570 511 50% 700 560 90% 830 633 Maximum 890 665 Sulfur, weight/percent. 1. 05 1. 19 Nitrogen, p.p.m.:

otal 2, 240 1, 470 Basic 822 275 Benzocoroncnes, p.p.rn Nil Nil Process conditions held relatively constant over the run were as follows:

The hydrolining catalyst employed was a presulided composite of 3.1 weight-percent nickel oxide and 14.7 percent M003 supported on a coprecipitated aluminasilica cogel containing 3.2 percent SiO2. The hydrocracking catalyst employed was a copelleted composite of 80 weight-percent of a magnesium-hydrogen Y zeolite (3 weight-percent MgO), and 20 weight-percent of an activated alumina binder, the composite containing 0.5 weightpercent of palladium added by ion-exchange.

Throughout the run, the various reactor effluents were sampled and analyzed for benzcoronenes by chromatographic adsorption of the respective samples on activated alumina, followed by elution with methylene chloride land ultraviolet absorption analysis of the eluents at 4120 A., the benzcoronene content being given as:

p.p.m. benzcoronenes=p4 412/ W where W is the weight of the sample and A4120 is the absorbence of the fraction at 4120 A.

In the hydrofiner, temperatures were gradually increased from 690 F. to 716 F. over the first 75 days, to maintain about 3 p.p.m. of basic organic nitrogen in the reactor effluent. During this period the benzcoronene content of the hydrofiner eflluent increased from about 7 p.p.m. to about 2l p.p.m. At this point hydrofining temperatures were raised at a more rapid rate over the next thirty days, from 716 to about 758 F. in order to reduce the effluent nitrogen content to below l p.p.m. This increased hydroiining severity resulted in a substantially immediate rise in the eluent benzcoronene content to about 40-50 p.p.m. However, the effluent from the first stage hydrocracker, which was operating at approximately the same temperature as the hydroiiner, continued throughout this period to show a benzcoronene content of about 10-20 p.p.m., indicating that the zeolite catalyst can convert benzcoronenes at these elevated temperatures and in the presence of ammonia (from the hydroner), down to this apparently equilibrium value.

During the first 88 days of the run, the second stage hydrocracker was operated in the absence of ammonia at temperatures increasing from about 520 to 566 F., and with total recycle of 400 F.+ bottoms thereto. Due to this total recycle, and the low temperatures employed, the benzcoronene content of the feed to the second stage reactor built up to about 120 p.p.m., which was reduced to only about p.p.m. in the effluent therefrom. It is thus aparent that the zeolite catalyst is relatively inelfective for converting benzcoronenes in the absence of ammonia and at low temperatures.

At this point, about 50 p.p.m. of ammonia (based on feed) was added to the second-stage recycle gas stream. Temperatures therein were concomitantly increased to about 615 F. over the next thirty days in order to maintain a desired conversion level. Within two days the influent benzcoronene content of the second stage feed dropped to about 25-35 p.p.m., while the eluent content dropped to the same equilibrium value of about 10-20 p.p.m. which was achieved in the first-stage hydrocracker.

It is thus apparent that the hydrofining type catalysts employed herein are prone to bring about synthesis of coronene ovalene compounds when utilized at high hydrocracking temperatures, while the zeolite type catalysts are effective for destroying such compounds when employed at relatively high temperatures in the presence of ammonia.

EXAMPLE VII Using the same hydroning catalyst and zeolite hydrocracking catalyst employed in Example VI, a hydrocracking operation according to the present invention was carried out with the feed passing first through the hydroning catalyst bed and then through the zeolite hydrocracking catalyst bed, with unconverted oil boiling above about 700 F. being recycled to the hydrofining catalyst |bed. The feed was a Kuwait vacuum gas oil having the following major characteristics:

Gravity, API

ASTM distillation D-1160" F.:

13 Product yields and properties were as follows:

TABLE 6I Yields, vol. percent of feed:

Butanes 9.3 C-C`6 17.7 Cfr-300 F. 17.1 300=700 F. 72.0 Product properties:

SOO-700 F. diesel- Gravity, API 41 Aniline Point, F. 162 Diesel Index 66 Pour Point, F. -40

Additional modifications and improvements utilizing the discoveries of `the present invention can readily be anticipated by those skilled in the art from the foregoing disclosure, and such modifications and improvements are intended to be included Within the scope and purview of the invention as defined in the following claims:

What is claimed is:

1. A process for the manufacture of a middle distillate product :boiling in the 3D0-700 F. range from an initial mineral oil feedstock boiling mainly between about 700 and 1000o F. with at least about 20 volume-percent boiling above 800 F., and also containing sulfur and nitrogen compounds, which comprises:

(l) subjecting said initial feedstock plus added hydrogen and recycle oil derived from step (4) to catalytic hydroning-hydrocracking in contact with a hydrofi-ning-hydrocracking catalyst at temperiatures between about 750-850 F., pressure above about 1,000 p.s.i.g., and at space velocities correlated with temperature to effect betwen about 40 70 volume-percent conversion of total feed to material boiling below about 700 F., said hydrofininghydrocracking catalyst comprising a Group VI-B metal oxide and/ or sulfide supported on an adsorbent mineral oxide carrier having a Cat-A cracking activity index -below about 25;

(2) subjecting total effluent from step (1) including Y ammonia, to catalytic hydrocracking at elevated pressures, temperatures between about 700-800 F., and at space velocities correlated with temperature and `ammonia concentration so as to achieve a substantial conversion of coronene-ovalene compounds while limiting overall conversion to less than thirty volume-percent of the unconverted oil from step (1), said catalytic hydrocracking being elected in the presence of a hydrocracking catalyst comprising essentially a Group VI-B and/ or Group VIII metal hydrogenating component supported on a zeolite cracking base having an ion exchange capacity greater than about 1.0 meq./gm., a crystal pore diameter greater than that corresponding to a Cat-A activity index of 40, and wherein the zeolitic cations are mainly hydrogen ions and/or polyvalent metal ions; (3) fractionating the efiiuent from step (2) to recover said middle distillate product and an unconverted oil boiling mainly above about 700 F.; and

(4) recycling said unconverted oil to step (1).

2. A process as defined in claim 1 wherein said hydrocracking catalyst comprises a Group VIII noble metal supported on a Y zeolite.

3. A process as defined in claim 1 wherein said hydrofning-hydrocracking catalyst comprises nickel and/or cobalt plus molybdenum and/or tungsten supported on a carrier which is essentially activated alumina.

4. A process as defined in claim 1 wherein the space velocity in step (2) is 2 to 4 times higher than the space velocity in step (l).

5. A process as defined in claim 1 wherein ammonia is added to the eiuent from step (l) so as to permit the use of higher temperatures in step (2) Without increasing the conversion level therein above about 30 volumepercent per pass'.

6. A process as defined in claim 1 wherein temperatures and pressures in step ,(1) are controlled and correlated so as to` achieve about 40-60 lvolume-percent conversion per pass to material boiling below about 700 F.

7. A process as defined in claim 1 wherein temperature, space velocity and ammonia concentration are controlled and correlated in step (2) so as to achieve about 10-25 volume-percent conversion therein to material boiling below about 7 00 F.

References Cited UNITED STATES PATENTS 2,120,295 16/ 1938 Pier et al. 208-108 3,287,252 ll/ 1966 Young 208-59 3,383,305 5/1968 Rogers et al. 208-89 DELBERT E. GANTZ, Primary Examiner A. RIMENS, Assistant Examiner U.S. Cl. X.R.

UNITED STATES PATENT OFFICE CERTIFICATE OF CGRRECTION Patent No. 3 554 898 Dated January 12 1971 Frederick C Wood et a1 Inventor(s) It is certified that error appears in the above-identified patent and that said Letters Patent are hereby corrected as shown below:

Column 14 line 5 after "than" insert about 8A. and

cracking activity greater than Signed and sealed this 11th day of May 1971 (SEAL) Attest:

EDWARD M.FLETCHER,JR. WILLIAM E SCHUYLER, Commissioner of Paten Attesting Officer llSCnMM-DC B03 

